Integral hydrogenation isomerization process

ABSTRACT

LIGHT GASOLINE FRACTIONS RICH IN N-HEXANE AND OPTIONALLY N-PENTANE, AND WHICH ALSO CONTAIN SMALL AMOUNTS OF BENZENE, ARE FIRST SUBJECTED TO HYDROGENTATION IN CONTACT WITH CERTAIN ZEOLITE-BASED PLATINUM-GROUP METAL CATALYST TO EFFECT SATURATION OF THE BENZENE CONTENT, AND TOTAL EFFLUENT FROM THE HYDROGENATION ZONE IS THEM SUBJECTED TO ISOMERIZATION IN CONTACT WITH A PLATINUM GROUP METAL SUPPORTED ON AN ALUMINA-CONTAINING BASE TO EFFECT ISOMERIZATION OF NORMAL PARAFFINS TO ISO-PARAFFINS. THE ISOMERIZATION CATALYST IS MAINTAINED IN AN ACTIVE STATE BY MAINTAINED A SMALL PROPORTION OF HYDROGEN CHLORIDE IN THE RECYCLE GAS.

Sept. 18, 1973 E. c. ATTANE ET AL 3,759,819

INTEGRAL HYDROGENATION-ISGMERIZATION PROCESS Filed June 30, 1971 ATTORNEY United States VPatent O 3,759,819 INTEGRAL HYDROGENATION-ISOMERIZATION PROCESS Edward C. Attane, Orange, David E. Mears, Fullerton, and Cloyd P. Reeg, Orange, Calif., assignors to Union Oil Company of California, Los Angeles, Calif.

Filed .lune 30, 1971, Ser. No. 158,447 Int. Cl. C07c 5/24; C10g 39/00 U.S. Cl. 208-57 10 Claims ABSTRACT F THE DISCLOSURE BACKGROUND AND SUMMARY OF THE HNVENTION With the advent of low-lead and lead-free gasoline there is a growing need in the petroleum industry for high octane hydrocarbons boiling in the gasoline range. This need cannot be satisfied entirely by increased catalytic reforming capacity which is useful primarily for upgrading the heavier gasoline fractions by aromatization. The light fraction boiling in the C-C6 range is composed mainly of paratlins which are not appreciably affected by catalytic reforming. To obtain maximum octane values in the light gasoline fractions it is necessary to resort to isomerization processes to convert n-pentane and n-hexane to iso-parafns, or to convert normal butane to isobutane for alkylation purposes. Low temperature isomeri zation of C5 and C6 parains is particularly desirable because the equilibrium at low temperatures favors isoparains. However, at these low temperatures (e.g., 200- 300 F.) conventional platinum-promoted, acidic isomerization catalysts are very sensitive `to minor amounts of benzene present in the feed. Even as little as about l0 parts per million of benzene greatly reduces catalytic activity at these low temperatures. Conventional proce dures for separating aromatics from paraflins are unecenomical when the benzene content must be reduced to less than 50 p.p.m. Distillation is impractical due to the hexane-benzene azeotrope, while solvent extraction requires a large number of stages. It has therefore been proposed in the past to carry out a preliminary, substantially complete hydrogenation of the benzene in such feeds before they contact the isomerization catalyst.

However, in previously proposed low-pressure isomerization processes, it has been suggested that the prehydrogenation step be carried out in a non-integral fashion, i.e., in a separate reactor with its own product separation and recycle gas facilities. Th'is adds greatly to capital and operating costs, but has in the past been considered necessary because of several interrelated economic factors. Firstly, the isomerization reaction does not require high hydrogen partial pressures, and it is hence desirable to econornize by carrying out this step in a relatively cheap, low-pressure reactor operating at, eg., 20G-500 p.s.i.g. But due to the relatively low activity of previously known hydrogenation catalysts, it has been a practical necessity to carry out the hydrogenation step at substantially higher hydrogen pressures of, e.g., 500 to 2,000 p.s.i.g. With a dual pressure system such as this, it is a ice practical necessity to operate in a non-integral fashion because to operate integrally would require depressuring the effluent from the hydrogenation reactor to the low pressure maintained in the isomerization reactor, and then repressure the recycle gas from the isomerization reactor to the higher pressure maintained in the hydrogenation reactor. The power consumption involved in such a system is prohibitive. It is hence necessary as a practical matter to operate an integral hydrogenation-isomerization system in a substantially isobaric fashion. But as noted, this has not heretofore been considered attractive due to the high pressure required in the hydrogenation zone when utilizing previously known hydrogenation catalysts, and the consequent, unnecessarily high pressure for which the isomerization step must be designed.

Another factor which has operated in the past to discourage the use of an integral system resides in the necessity of maintaining a volatile halide such as HC1 in the system in order to activate the isomerization catalyst. In an integral system, this volatile halide must be recirculated in the recycle gas through the hydrogenation zone and the isomerization zone. Hydrogen chloride has been found to substantially reduce the hydrogenation activity of previously known hydrogenation catalysts. But the hydrogenation catalysts employed herein are not only extremely active in the absence of "HC1, but they lose only a small proportion of this activity when hydrogen chloride is added to the system. It will thus be seen that the key to the success of our low-pressure, integral hydrogenation-isomerization process resides principally in the nature of the hydrogenation catalyst employed.

The best catalysts developed to date for hydrogenation of aromatic hydrocarbons consist of platinum impregnated onto various alumina or alumina-silica supports. Conventional platinum-alumina reforming catalysts are sometimes used. We have now discovered that the Group VIH noble metals, particularly platinum or palladium, when supported upon certain acidic, crystalline zeolite bases, exhibit extraordinarily high hydrogenation activity such that effective hydrogenation of benzene can occur at temperature of 250-500 F., pressures of 200 to 750 p.s.i.g., at space velocities of 5 to 40, and that these conditions can be correlated so that little or no cracking occurs. Further, it has been discovered that the necessary hydrogen chloride for activating the isomerization catalyst can be passed through the hydrogenation catalyst bed under the described hydrogenation conditions without substantially reducing the activity of the catalyst or destroying its crystal structure.

DETAILED4 DESCRITION (A) Description of process The process details will be described in conjunction with the accompanying drawing, which is a ow sheet illustrating the process in some of its preferred modifications. The initial feedstock to be isomerized may cornprise any hydrocarbon fraction containing n-hexane in higher-than-equilibrium proportions with respect to its other isomers. The feed should be substantially free of non-hydrocarbon impurities such as sulfur, nitrogen, oxygen and the like, and may also comprise a lower boiling fraction rich in normal pentane. Benzene is present normally in amounts ranging between about 0.1 and 6 percent by volume in feedstocks derived from most refinery sources. In the modification illustrated, the intial feed is brought in via line 2 and passed through product heat exchangers 4 and 6. The partially heated feedstock in line 8` is then mixed with recycle and makeup hydrogen from line 10 and makeup organic chloride from line 12. The resulting mixture is then heated to the desired hydrogenation temperature in final heater 14 and passed into hydrogenation reactor 16, where benzene hydrogenation proceeds in the presence of the unique hydrogenation catalyst to be described more in detail hereinafter, and under the following hydrogenation conditions:

HYD RO GENATION CONDITIONS It will be understood that the above conditions should be correlated so as to achieve the desired degree of hydrogenation of benzene. Normally, it is essential to reduce the benzene content to below about 50 p.p.m., and preferably below p.p.m. Por these purposes, it will be understood that temperatures in the high ranges prescribed above will ordinarily be employed in conjunction with the higher space velocity ranges, and the lower temperatures will be employed in conjunction with the lower space velocity ranges. The contacting may be continued with periodic temperature increases to compensate for catalyst deactivation until undesirable cracking begins to occur. At this point the catalyst may be regenerated.

It should be noted that the organic chloride added via line 12 is decomposed in reactor 16 to form hydrogen ity. It is therefore preferred to add the makeup chloride itself is somewhat detrimental to the hydrogenation activity of the catalyst in reactor 16, organic chlorides such as carbon tetrachloride, chloroform, methylene dichloride and the like appear to enhance the hydrogenation activity. It is therefore preferred to add the makeup chloride to the hydrogenation reactor, rather than to the isomerization reactor. It should be noted by way of precaution that all streams fed to reactor 16 should be substantially anhydrous (preferably less than 10 p.p.m. of water), to prevent acid attack upon the catalyst, as well as corrosion of the reactor and its internals.

Total effluent from reactor 16 is then transferred via heat exchanger 6 into isomerization reactor 18. In exchanger 6, the efliuent is partially cooled to the desired isomerization temperature, and any additional cooling required may be obtained by diverting a portion of the cool recycle gas via line 20 into line 22. Reactor 18 is filled with a suitable isomerization catalyst described more particularly hereinafter, and isomerization occurs therein under the following general conditions:

ISOME RIZATION CONDITIONS Broad Preferred range range Temperature, F 175-400 20D-300 Pressure, p.s.i.g 200-1, 000 250-450 HSV 0. 5-10 1-5 Hz/oil, M s.e.. 0. 5- 1-5 tinued until cracking temperatures are reached, whereupon the catalyst may be regenerated. To maintain the desired catalytic activity in reactor 18, it is necessary to maintain about 0.01 to 5 percent by Volume of hydrogen chloride in the gas system. Preferred hydrogen chloride ratios range between about 0.1 to 1.0 volume percent.

Effluent from reactor 18 is then cooled in heat exchanger 4 as previously described, and then passed via line 24 into high pressure gas-liquid separator 26, from which recycle hydrogen and hydrogen chloride are withdrawn via line 28 and recycled as previously described. Liquid product is withdrawn via line 30. In many cases, sufficient octane improvement of the feed can be obtained in a oncethrough operation, in which case the desired product is withdrawn from the process via line 32 and sent to blending or storage facilities not shown. However, if further octane improvement is desired, all or a portion of the liquid product from separator 26 may be diverted via line 34 to fractionating column 36. The modification illustrated assumes a substantially pentane-free C6 feedstock. In this case the more highly branched hexane isomers such as neohexane (2,2 dimethyl butane), and 2,3 dimethylbutane are withdrawn as overhead via line 38. The unconverted normal hexane, along with the methyl pentanes if desired, may then be recycled via line 40 to reactor 18, so that eventually a substantially total conversion to the more highly branched isoparains is obtained.

In case the feedstock is one containing a substantial C5 fraction in addition to the C6 fraction, an additional fracseparate may be employed ahead of fractionator 36 to separate `out the neopentane overhead, a normal pentaneisopentane sidecut which is recycled to reactor 18, and the C6 bottoms product which is then sent to fractionator 36 as previously described. The choice of these various recycle procedures depends largely upon economic factors which forms no essential part of the invention.

(B) Hydrogenation catalysts The hydrogenation catalysts contemplated herein comprise as the hydrogenating component one or more of the Group VIIvI noble metals in amounts of about 0.05-4 percent by weight. Specifically included are the metals platinum, palladium, rutheniurn, rhodium, iridiurn and osmium, with platinum being preferred. These hydrogenating metals are supported on a crystalline aluminosilicate zeolite base having relatively uniform pore diameters between about 6 and 15 A., and which has been converted either partially or completely to a hydrogen (protonated) and/or decationized form. Specifically, it is preferred that the zeolite base be at least about 20 percent, and preferably at least about 35 percent metal-cationdeficient as a result of conversion to the hydrogen and/or decationized forms. These zeolites are well known for their cracking activity, and while cracking activity is generally regarded as being substantially independent of hydrogenation activity, it has been found that the hydrogenation activity of Group VIII noble metals is greatly enhanced when supported upon these acidic, metal-cationdefcient zeolites. This is aptly illustrated by a series of low temperature hydrogenations carried out with several different palladium-zeolite catalysts, wherein naphthalene and tetraline feedstocks were subjected to hydrogenation in a stirred autoclave:

TABLE 1 Moles Hz consumed Hydrogen per hr. per gm. of Pd Temp pressure, lieed Catalyst p.s.i.g. NT l T-D Tetralin 0.53% Pd on hydrogen Y zeolite 200 167 1. 9 D0... 1% Pd on Magnesium Y zeolite 200 167 0. 49 o 1% Pd on sodium Y zeolite 200 167 0.28 Naphthalen 1% on nuxed hydrogen-magnesium Y 250 172 1.5

zee i e.

Do 1% Pd on magnesium Y zeolite 250 172 0. 63 D0 1% Pd on sodium Y zeolite 250 172 0.56

1 Naphthalene to tetralin. .Tetralin to Decalin.

Thus, palladium deposited on hydrogen zeolites appears to exhibit a hydrogenation activity about 3 to 7 times that of palladium deposited upon sodium zeolites or magnesium zeolites. Hence, a preferred catalyst is one wherein a substantial amount (e.g., 50-90%) of the total exchange capacity has been protonated and/or decationized, and a most preferred catalyst is platinum on a hydrogen and/ or decationzed Y zeolite. Other zeolites which may be utilized include for example the X, L and Q forms, and synthetic mordenite which has been subjected to caustic or acid leaching to increase the pore size. The preferred zeolites have a SiOz/AlzOs mole-ratio between about 3.5 and 7.0.

An especially preferred hydrogenation catalyst for use herein comprises a stabilized Y zeolite base which, after the addition of noble metal thereto, has been calcined at temperatures above about 1000 F., for a sutlicient time to give a weight loss on ignition (LOI) at 2000" F. of less than about 4 weight-percent, preferably between about 0.5 and 3 weight-percent. The final, high-temperature calcination is found in many cases to more than double the activity of the resulting catalyst. It should be noted however that in order to carry out this high temperature calcination it is necessary to start with a thermally stable form of the Y zeolite. Thermal stability up to temperatures of about 1500-1800" F. may be obtained by two general procedures, or a combination of both. Firstly, after replacing at least about 80%, and preferably at least about 90% of the original sodium ions with ammonium ions, the resulting ammonium zeolite may be partially back-exchanged with solutions of polyvalent metal salts to replace from about -80%, preferably 20-50% of the ammonium ions with an ionic equivalent of polyvalent metal ions. The resulting mixed polyvalent metal ammonium zeolites may then be dried and calcined at temperatures of, c g., 700 to 1500 F. to produce a stable polyvalent metal-hydrogen zeolite. Suitable polyvalent metal ions include those of the metals of Groups II-A,l III-A, II-B, VII-B and VIII, as well as the rare earth metals. Preferred polyvalent metal cations comprise magnesium, calcium, nickel, zinc and the rare earth metals, eg., cerium, lanthanum, praseodyminum, neodymium, samarium, etc.

To achieve maximum activity and stability however, it is normally preferable to convert substantially all of the exchange sites to decationized and/ or hydrogen ion sites. Zeolite bases of this nature are prepared by exhaustive exchange of the original sodium ions with ammonium ions, followed by a special hydrothermal treatment to effect stabilization. Maximum activity is obtained from zeolite bases wherein the sodium content has been reduced to less than 1:5 weight-percent NazO, preferably less than 0.5 percent. The hydrothermal stabilization treatment consists essentially in heating the ammonium ion exchanged zeolite in a steam atmosphere at temperatures of about 900-J1500 for times ranging between about 30 minutes and 12 hours or more. A single such stabilization treatment produces a very acceptable zeolite base, but even better results are obtained by subjecting the once-stabilized zeolite to further ammonium ionexchange to reduce the sodium content even further, and then subjecting it to a second hydrothermal treatment of the same nature. The rst hydrothermal treatment effects a redistribution of diiicultly exchangeable sodium ions remaining after the initial ammonium ion exchange, such that during the second exchange the nal sodium content can be reduced to very low levels of, e.g., 0.1 weightpercent NaZO or less.

The mechanism of hydrothermal stabilization is not completely understood, but appears to involve removal of aluminum atoms from the anionic crystal lattice structure, with a resultant increase in the SiO2/A12O3 moleratio of the structure. Removal of aluminum atoms also brings about a shrinkage in the unit cell size of the zeolite crystals, and this reduction in cell size is a fairly reliable indicia of the stability achived. The smaller the unit cell size the greater the stability as a general rule. The unstabilized hydrogen Y zeolite normally displays a unit cell size of about 24.65v A. Some degree of stabilization is indicated by a unit cell size of 24.60 A., but for maximum stability the cell size should be reduced to between about 24.20 and 24.55 A., preferably 24.30- 24.50 A. The SiO2/Al203 mole ratio of such zeolites ranges Ibetween about 5.0 and 8.0.

The noble metal is added to the stabilized zeolite base by conventional procedures such as impregnation with aqueous solutions of the desired salts, or preferably by ion exchange with aqueous solutions of the noble metal compounds in which the noble metal appears in the cation. Suitable compounds for ion exchange from aqueous solutions include primarily the complex ammino compounds such as platinous tetramminochloride, platinous tetraamminohydroxide, and the like.

The zeolite base may be formed. into the shape desired for the final catalyst either before or after the addition of noble metal. Typical procedures consist of compressing the powdered material into pellets in a tableting machine, subjecting the moistened powder to prilling, extrusion or the like. Prior to these operations it is normally desirable to admix with the zeolite base about 10-30 weight-percent of a suitable inorganic binder such as clays, alumina gel, alumina-silica cogels, or the like.

(C) Isomerization catalyst In broad aspect, any isomerization catalyst may be employed herein which can be activated by hydrogen chloride or organic chlorides so as to provide a substantial n-paraflin isomerization activity in the temperature range of about 20G-350 F. Catalysts of this general nature are well known in the art, and hence need not be described in detail. In general, the most practical catalysts are those composed of about 0.05 to 2 weightpercent of a Group VIII noble metal supported on a porous carrier which comprises a substantial proportion of alumina. 'Ihe preferred Group VIII metal is platinum, and preferred support bases comprise alumina, or alumina-silica cogels. An especially preferred support comprises eta alumina, or mixtures of eta and gamma alumina. Preferably, such catalysts are preactivated by reaction with a suitable organic chloride at temperature between about 300 and 650 F. Preferred organic chlorides comprise aliphatic chlorohydrocarbons having an atomic ratio of chlorine to carbon of at least 2 to 1. Suitable isomerization catalysts are described more in detail in U.S. Pat. Nos. 3,287,439, 2,999,074, 3,551,516 and 3,242,228.

The following examples are cited to illustrate the critical aspects of the invention, but are not to be construed as limiting in scope:

EXAMPLE I-HYDROG-ENATION CATALYST PREPARATIONS A preferred hydrogenation catalyst (designated catalyst A) of the present invention was prepared as follows:

An ammonium Y zeolite containing 1.6 weight-percent NaZO was first calcined in a closed vessel (to retain the generated steam atmosphere) at 1012 F. in a preheated oven, then exchanged twice with a 20% ammonium sulfate solution, and finally recalcined in the same manner at 1300o F. for three hours. The final sodium content was 0.093 weight-percent as Nazi). Platinum was then exchanged into the stabilized zeolite from an aqueous solution of platinic hexamminochloride to provide 0.5 Weight-percent platinum based on the zeolite. The resulting material was then mulled with peptized alpha alumina monohydrate (Boehmite) and then extruded through a 1i6-inch die. The final catalyst contained weight-percent Y zeolite and 0.39 weight-percent Pt. The unit cell size was 24.359 A. Half of this batch of catalyst was then subjected to a final calcination at 1200 P. in owing dry air for one and one-half hours to give a product having an L.O.I. of 1.87 weightpercent.

A second portion of the platinum-exchanged zeolite was subjected to a final calcination for two hours at 1020 F. in flowing dry air to give a catalyst (designated B) having a relatively higher ratio of Bronsted/Lewis acidity, as indicated by an L.O.'I. of 3.08 weight-percent.

Comparison catalyst C was a commercial reforming catalyst comprising 0.38 weight-percent of Pt supported on a mixed eta-gamma alumina base.

EXAMPLE II-ACTIVITY TESTING The catalysts of Example I were tested for hydrogenation activity, using as the feedstock a hydrocracked gasoline fraction consisting primarily of petanes and hexanes, and having a gravity of 72.1 API and a benzene concentration of 4.7 weight-percent (47,000 ppm.) The feedstock was processed by passing it through beds of the respective catalysts at 360 F., 380 p.s.i.g. with 2600 s.c.f./b. of hydrogen. In -view of the drastic difference in activities of the respective catalysts it was necessary to carry out the various runs at different space velocities, and then calculate first order rate constants, k, from the equation:

Heurs on stream hr.-l benzene, p.p.m. stent, k

It is readily apparent from the foregoing that catalyst A (calcined at 1200 F.) is more than twice as active as catalyst B (calcined at 1020 F.), and is about 20 times as active as the more conventional hydrogenation catalyst C. Cracking was negligible in all the runs.

The foregoing hydrogenation test runs were carried out in the absence of added hydrogen chloride. In systems containing about 0.5 volume-percent of hydrogen chloride in the recycle gas, the rate constants for catalysts A and B are reduced only about 30 to 40 percent, while the rate constant for catalyst C is reduced by about 50-60 percent. The superiority of catalysts A and B for use in an integral hydrogenation-isomerization process employing a recycle gas containing hydrogen chloride is thus readily apparent.

EXAMPLE III Hydrogenation zone IIg/oil ratio, s.c.l'./b 2, 000 2, 000 LHSV, lum-1 21 3 Reactor pressure, p.s.i.g 400 375 Average bed temperature, F. 415 250 HC1 in recycle gas, VOL-percent 0.6 0, 5

Isomerization Under these operating conditions, the characteristics of the feed and product are as follows:

Product Greater octane improvements are obtained with a feed containing a higher percentage of normal isomers, or if the reaction is continued to equilibrium.

To obtain these same results using catalyst C in both the hydrogenation and isomerization zones, other process conditions being the same, it is found that a space velocity of about l is required in the hydrogenation zone, which means that at this particular hydrogenation temperature, about 20 times as much hydrogenation catalyst is required to achieve the same results.

It is not intended that the invention should be limited to non-essential details described above; the following claims and their obvious equivalents are intended to de,- fine the true scope of the invention:

We claim:

1. A process for isomerizing a hydrocarbon feedstock containing an isomerizable proportion of n-hexane and a minor proportion of benzene, which comprises:

(l) subjecting said feedstock to hydrogenation with added hydrogen in a hydrogenation zone maintained at a temperature between about 250 and 600 F. and a pressure between about 200 and 1,000 p.s.i.g., in contact with a catalyst comprising a Group VIII noble metal supported on a crystalline aluminosilicate zeolite having crystal pore diameters between about 6 and 15 A., at least about 20% of the ion-exchange capacity of said zeolite being decationized and/or protonated, to thereby effect a substantially complete hydrogenation of the benzene in said feedstock;

(2) subjecting total effluent from said hydrogenation, without intervening gas separation to isomerization in an isomerization zone maintained at a temperature between about and 400 F. and a pressure between about 200 and 1,000 p.s.i.g., in Contact with a catalyst comprising a Group VLII noble metal supported on a porous, partially chlorided alumina-containing carrier;

(3) separating the effluent from Step (2) into an isoparaffin-rich product stream, and a hydrogen-rich recycle gas;

(4) recycling at least a portion of said recycle gas t0 Step (1); and

(5) maintaining in said recycled recycle gas a concentration of hydrogen chloride between about 0.01-5 volume percent.

2. A process as dened in claim 1 wherein said feed stock also contains n-pentane in isomerizable proportions.

3. A process as defined in claim 1 wherein the catalyst employed in Step (1) is platinum supported on a stabilized Y zeolite.

4. A process as dened in claim 1 wherein the catalyst employed in Step (1) is platinum supported on a steamstabilized, decationized and/ or protonated Y zeolite, which catalyst has been calcined at a temperature above about 1000" F. for a sufficient time to reduce its loss-on-ignition to below about 4 weight-percent.

5. A process as defined in claim 1 wherein the catalyst employed in Step (2) is platinum supported on an etaalumina base, which catalyst has been pre-activated by reaction with an aliphatic chloride h'aving an atomic ratio of chlorine to carbon of at least 2 to l.

'6. A process as defined in claim 1 wherein sufficient of an aliphatic chloride having an atomic ratio of chlorine to carbon of at least 2 to 1 is added to the recycled recycle gas in Step (5) to maintain the stated proportion of hydrogen chloride therein.

7. A process as dened in claim 1 wherein the catalyst employed in Step (1) is platinum supported on a steamstabilized, decationized and/or protonated Y zeolite, which catalyst has been calcined at a temperature above about 1000 F. for a suflicient time to reduce its loss-onignition to below about 4 weight-percent, and wherein the catalyst employed in Step I(2) is platinum supported on an eta-alumina base, which catalyst has been pre-activated by reaction with an aliphatic chloride having an atomic ratio of chlorine to carbon of at least 2 to l.

8. A process for isomerizing a hydrocarbon feedstock containing an isomerizable proportion of n-hexane and a minor proportion of benzene, which comprises:

(1) subjecting said feedstock to hydrogenation with added hydrogen in a hydrogenation zone maintained at a temperature between about 350 and 500 F. and a pressure between about 250 and 500 p.s.i.g., in Contact with a catalyst comprising platinum supported on a stabilized Y zeolite wherein at least about 20 percent of the ion exchange capacity thereof is decationized and/or protonated, to thereby reduce the benzene content of the feedstock to below about ppm.;

(2) subjecting total eiuent from said hydrogenation zone, without intervening gas separation, to isomerization in an isomerization zone maintained at a ternperature between about 200 and 300 F. and a pressure between about 250 and 450 p.s.i.g., in contact with a catalyst comprising platinum supported on an eta-alumina base, which catalyst has been pre-activated by reaction with an aliphatic chloride having an atomic ratio of chlorine to carbon of at least 2 to 1;

(3) separating the efiluent from Step (2) into an isoparailin-rich product stream, and a hydrogen-rich recycle gas;

(4) recycling at least a portion of said recycle gas of Step (1); and

(5) adding to said recycled recycle gas sufficient of an aliphatic chloride havingy a chlorine to carbon atomic ratio of at least 2 to l, to maintain therein a concentration of hydrogen chloride between about 0.01 and 5 volume-percent- 9. A process as defined in claim 8 wherein said feedstock also contains n-pentane in isomerizable proportions.

10. A process as defined in claim 8 wherein the catalyst employed in Step (1) is platinum supported on a steam-stabilized, decationized and/or protonated Y zeolite, which catalyst has been cacined at a temperature above about 1000 F. for a suticient time to reduce its lossonignition to below about 4 weight-percent.

References Cited UNITED STATES PATENTS 2,493,499 1/1950 Perry 208--57 3,287,439 11/1966 Suggitt et al 260-683.68 3,527,695 9/1970 Lawrence et al 260--667 3,197,398 7/1965 Young 260-667 HERBERT LEVINE, Primary Examiner 

